Catalytic conversion of hydrocarbons



July 18, 1950 P. c. KEITH ETAL CATALYTIC CONVERSION OF HYDROCARBONS 3 Sheets-Sheet 1 Filed April 24, 1941 PERC/YHL c. KEITH- JOSEPH w JEWELL INVENTORS av fii/ ATTORNEYS July 18, 1950 v P. c. KEITH HAL 5 3 CATALYTIC CONVERSION OF HYDROCARBONS JOSEPH n1, JEWELL.

F/G- 2 f ATTORNEYS Jufly 3, E950 'P. Q. KEIITH m.

CATALYTIC CONVERSION 01-" mrmocmoms 3 Sheets-Sham 3 Filed April 24, 1941 Fatented July 1950 CATALYTIC CONVERSION OF mmnocannons Percival 0. Keith, P'eapaclr', and Joseph w. Jewell, Summit, N. .l.,' auignors to. The M. W. Kellogg Com ny, Delaware Jersey City, N.-J., a corporation of Application April 524, 1941 Serial No. 390,164

' 3 Claim 3 (CL 196-52) in process and apparatus for efiecting catalytic conversions. is directed particularly to an improved continuous process of converting hydrocarbons by treat In its specific aspects the invention The present invention relates to imprbv fi I process is subject to certain limitations which are fobviated by the present process. The primary object of the present invention is the provision of aprocess wherein the catalyst is introduced into ment over catalytic materials which become spent or deactivated during the conversion, and which accordingly require periodic regeneration treatment to fit them for reuse in the conversion step. The catalytic conversion of high boiling hydrocarbons such as petroleum gas oil and the like into low boiling hydrocarbons within the gasoline boiling range is an example of the latter type of conversion reaction of outstanding importance.

It has been proposed heretofore to catalytically convert high boiling hydrocarbons such as gas oil to low boiling hydrocarbons within the gasoline boiling range by passing vapors of the high boiling hydrocarbons under suitable reaction conditions in contact with a stationary bed of a cracking catalyst disposed in a catalyst chamber. Pursuant to such processes, after the activity of the catalyst is decreased by reason of the formation of a carbonaceous deposit thereon to an extent where regeneration is necesary or desirable, the activity of the catalyst is restored by stopping the fiow of oil vapor to the chamber and passing an oxygen-containin gas into the chamber in contact with the spent catalyst, thereby regenerating it in situ by combustion of the carbonaceous deposit. Although such processes are commercially practicable they are subject to a number of inherent limitations and disadvantages which are eliminated by the present invention. Among these are the intermittent nature of the operation, variations in product quality and quantity during the reaction period and difllculty in temperature control, particularly in the regeneration operation.

A process whereby various of the objectionable limitations and-disadvantages of the stationary bed or intermittent type of operation are eliminated is described in U. S. Patent No. 2,350,730 to Keith et al., dated June 6, 1944. In accordance with the Keith et al. process continuous operation and uniformity in yield and quality of products in the practice of catalytic reactions such catalytic cracking and the like, is obtained by a procedure wherein the vapors of the oil undergoing conversion are passed in contact with a compact mass or moving bed of catalytic material confined in a reaction zone, and the average catalytic activity of the catalytic mass is maintained substantially constant by adding active catalyst thereto and removing spent catalyst therefrom without interruption of the" flow of oil vapor through the reaction zone.

In the process disclosed by said Keith et al. application the catalyst is employed in granular condition and the movement of the catalyst through the conversion zone is eilected by gravity flow at regulated rates controlled by suitable mechanical discharging means. Accordingly, the

v the conversion and regeneration zones in finely divided or powdered condition and at the same time certain of the variable operating conditions are so controlled as. to maintain a suitable concentration of the catalyst in said zones for the desired conversion and regeneration effects in an apparatus of feasible dimensions. The various features and advantages of the process will be evident from the following detailed description thereof given in connection with the appended drawings wherein:

Fig. 1 illustrates diagrammatically a suitable complete process flow and arrangement of apparatus for use in the practice of the invention.

Fig. 2 is an elevation view of the reactor or conversion chamber and the catalyst regeneration chamber or regenerator, andillustrates the details of these elements and their interconnection.

Fig. 3 is a view similar to Fig. 2 of a modified form of reactor and regenerator.

Fig. 4 is plan view illustrating the details of the balile device.

The steps of catalytic conversion and regeneration of the catalyst as preferably practiced pursuant to this invention may be more readily understood by reference to Fig. 2. As indicated thereon, a finely divided or powdered active catalyst is introduced through the outlet l of a catalyst standpipe 3| into a stream of the feed vapors traveling at a relatively high velocity through the reactor inlet line 2. Both the catalyst and vapors preferably are heated prior to their mixture in their line 2 to an elevated temperature suitable for the subsequent conversion. Vaporized feed may be supplied to line 2 by a transfer line 3 leading from a heater 4 or other suitable source of vaporized feed stock. Catalyst thus introduced is picked up by the vapors and carried therewith through line 2 into a conical inlet 5 in the lower part of reactor 6. Reactor 6 is a vessel, in the form of a cylinder or other suitable shape, having a relatively great cross-sectional area compared to the cross-sectional area of the vapor inletline 2, and these relative proportions cause a corresponding reduction in the velocity of the vapors after their passage from inlet line 2 into the reactor 6. The velocity of the vapors in reactor 6 is preferably maintained within such limits as to produce a concentrated or dense phase of the catalyst which state may conveniently be designated as] fluidized." This fluidized condition, 'in general, is characterized by the relatively high concentration of catalyst measured interm's of the quantity of catalyst per unit volume' of reactor space, and by the maintained low velocity of reactant vapors through the reactor.

The reactant vapors travel upwardly through the reactor in contact with the fluidized catalyst and during this period of contact undergo the desired conversion. Operating conditions in the reactor determined by variables such as the dimensions of the reactor, and the temperatures and rates at which reactant vapors and catalyst are supplied thereto, are maintained within such limits as to bring about the desired quality and extent of conversion as described in detail hereinafter.

The vaporous reaction products are withdrawn from the upper part of the reactor through a suitable outlet pipe or cone 1. It is contemplated that the used or spent catalyst may be separated from the vaporous conversion products by any one of several novel methods. According to that illustrated in Fig. 2, the vapors pass from the upper portionof the reactor into the conical outlet I of decreasing cross-sectional area wherein their velocity is progressively increased and then into an outlet pipe 8 of relatively restricted crosssectional area compared to that of the reactor. The vaporous conversion products mixed with spent catalyst exit through the outlet pipe at a relatively high velocity into a settling chamber or collecting hopper 9 of such cross-sectional area. that the velocity of the vapors therein is preferably of about the same order of magnitude but may be more or less than the vapor velocity in reactor 6. A bame l0, shown in plan in Fig. 4, is preferably interposed directly in the path of the vapor mixture exiting from pipe 8 whereby the mixture is directed laterally and downwardly thus functioning to propel catalyst particles present in the mixture out of the path of the vapor flow into a quiescent collecting zone defined by the inner walls of the outlet cone I and outlet pipe 8 and the lower inner walls of the settling or collecting hopper 9. Catalyst thus separated is withdrawn through suitable means such as catalyst standpipe H opening into the lower part of the collecting zone. A quantity of catalyst is preferably left at all times in said zone same order of magnitude as those utilized carry the particles through reactor 6.

Residual catalyst left in the vaporous conversion products exiting through line i3 to separated in a suitable recovery system described in detail hereinafter in connection with the Fig. 1, and may be returned to the spent catalyst sep arated in chamber 9 through line I4.

A suitable stripping medium such as steam is introduced through a line l5 having suitable fluid distributing means I6, in the bottom of the mass of catalyst in the collecting zone to strip or displace hydrocarbon vapors absorbed thereon or entrained therewith and to maintain the mass in an aerated flowable condition. While only one such line ii for introduction of the stripping medium is shown, it is to be understood that any suitable number may be employed and be so distributed throughout the collecting chamber as to assure the required stripping and aerating effects. The quantity of stripping steam is preferably such that its velocity in the collecting zone is low, that is, of the order of about 0.1 to 0.3 foot per second. The stripping medium and stripped oil vapors pass out of chamber 9 overhead through line i3 together with the vaporous conversion products.

' As illustrative of the operating conditions preferably maintained in the reaction zone reference is made to the data tabulated in the following Tables 1-A and 1B, wherein Table 1A represents conditions suitable for a large scale commercial unit for a given type of charging stock and capacity, and Table 1B, contains data obtained in the operation of aunit on a pilot plant or laboratory scale.

Feed weight ratio of catalyst to oil 5 to maintain a level of catalyst therein at a sub- Reactor temperature, inlet cone 5, F 933 stantial distance above the spent catalyst out- Reactor temperature, outlet cone 7, F.-- 900 l p n as indicated by dotted line l2. Vapor- Reactor pressure, inlet cone 5, 1bs./sq. in 13.0 ous conversion products are withdrawn from the Reactor pressure, outlet cone 7, upper part of the collecting hopper through line lbs /sq 11L 9.6 [3 mixed with a relatively small portion of the Vapor velocity, inlet, ft./sec 1.48 catalyst originally present in the mixture pass- Vapor velocity, outlet, ft./sec 2.45 ing through pipe 8. The effectiveness of the Ratio of weight of oil fed/hour to weight separation in chamber 9, as demonstrated hereof catalyst in reactor (w./hr./w.) 2.6 after by specific examples. is much greater than 011 vap r, on a m seconds 13.6 that predictable from theoretical conclusions Catalyst time, seconds 290 based upon a consideration of such known faccatalyst concentration, lbsjcu, it tors as the particle sizes and settling rate of (a) Inlet line 2 0.98 the catalytic material employed and the fact that (b) Reactor 18.0 the vapor velocity in chamber 0 may be of the (c) Reactor outlet, line 8 0.6

TABLE 1-B Run Number 5204 5325.4 5320.1 5205 533111 533015 on Feed, Liters/Hr.(3l.l .41 1 gas oil) 11.1 15. 0 15. 2 19. s 15.1 15.1 Steam Feed, Me] Per Cent 52.1 01. 3 00. 4 58.3 58. 2 02. 2 Reactor Dimensions:

5) Ht.-F 12 12 12 12 12 12 Dla.-ln 25 2.5 2.5 2.5 2.5 25 Weight Feed Ratio of Catalyst to Oil..- 4. 0 3-1 3. 6 2. 1 Reactor Temperature, Avem iesq 900 900 895 900 985 085 Reactor Pressure, Average, L .lSq. In. 10. 5 10.5 9.6 9.5 11.5 10.5 Vapor Velocity, Inlet, Fig/Sen 1.13 l. 48 l- 54 2. 03 1. 53 l. Vapor Velocity, Outlet, FtJSee l- 82 2. 22 2. 24 2. 62 2. 58 2. 87 Ratio of Weight of Oil Fed/H1. mWt. oi-Catalyst in Reactor WJHrJW- 3. 8 0. 4 8. 0 l3. 5 8. 3 7. 4 Oil Vapor, contact time, seconds 8 6. 5 (i 4 5. 2 5. 8 5. 2 Catalyst Time, is 240 350 ms 114 Catal t Concentration, Lbs/Cu. Ft: 1 5 m... 0.15 0.58 0. 05 0.31 0. 42 0.11. Yields I 11.0 12.3 10.3 1.2 10.3 11.5 mi Gasoline, v51. Per Cent... 41.1 43. 3 43. 0 s1. 4 43. 0 4113 2;; Excess Butane- 4.0 2.6 3.4 1.2 5.5 7.1 0 Cycle on 41.0 54.3 52.5 sac 40.2 42.0 a) Gas, Wt. Per Cent. 4.1 3.2 3.2 1.8 5.1 0.1 1! Carbon 5.1 3.5 4.4 1.2 3.5 4.4 Gasoline Octane No.- 79. 1 70. 0 80. 7 81. 7 U. 0

sorbed or entrained with the regenerated catalyst and maintain the separated catalyst in an aerated suitable provision is madefor any difierence in pressure between these zones. A somewhat higher pressure is normally preferably maintained in the bottom or inlet portion of the regeneration zone than the pressure maintained in the collecting zone, and a head of suitably aerated or fluidized catalyst is preferably maintained in theoutlet standpipe ll of a sufficient magnitude to balance or exceed this diflerential pressure. For this purpose spent catalyst flowing through standpipe H is maintained in a condition in which it has the flow characteristics of a liquid by introducing in suitably regulated amounts an aerating medium such as steam through lines H at the bottom of and at other suitably spaced points along the length of pipe I From the bottom of standpipe H spent catalyst is fed under the influence of the pressure head maintained therein and the pressure head provided by the mass of aerated catalyst in chamber 9 through a suitable feeding means such as a slide valve l8 into the regenerator inlet line l9.

Spent catalyst thus introduced is mixed with air or other suitable carrying medium such as steam introduced into pipe l9 by line In case air is employed, the quantity introduced is so controlled that the combustion of the spent catalyst in line I9 is not sufiicient to raise the temperature of the catalyst beyond the maximum safe regeneration temperature.

The mixture of hot spent catalyst and carrying medium flows through line l9 into an inlet cone 20 at the bottom of the regenerator 2| where it I meets and mixes with a stream of relatively cool recycled regenerated catalyst and air from cooler 22 and passes therewith upwardly through the regeneration chamber 2|. Operating conditions in the regeneration chamber or zone 2| are preferably maintained to provide a condition-similar to that maintained in the reaction zone with respect to a fluidized condition of the catalyst. This condition similar to that maintained in the reactor is characterized by the relatively large concentration of catalyst and low vapor velocities maintained in the regeneration zone. During the course of the travel of the spent catalyst upwardly through regeneration chamber 2| combustion of the carbonaceous deposit thereon is effected to the required extent at an elevated temperature maintained below the safe maximum regeneration temperature by means of the cooled recycled catalyst.

Gaseous regeneration products (flue gas) and regenerated catalyst exit from the upper part of the regenerator through an outlet 23 and into a separator 24 similar in design and mode of operation to separator 9 described in connection with the reactor. The major portion of the regenerated catalyst is separated and collected in a collecting zone at the bottom portion of the regenerated catalyst collecting hopper or separator 24 and the gaseous combustion products together with a relatively small amount of regenerated catalyst pass out overhead from chamber 24 through line 25 to a suitable recovery system such as that illustrated by and hereinafter described in connection with Fig. 1. Catalyst recovered from the vapors in line 25 may suitably be returned to hopper 24 through line 26.

Suitable means 21 and 28, similar to pipe l5 and distributor l6, are provided in the lower portion of hopper 24 to introduce a suitable medium such as steam to strip and displace flue gas abflowable state. As in the case or the reactor a level of separated catalyst indicated by dotted line 29 is preferably maintained at-a substantial distance above the catalyst outlet lines.

Regenerated catalyst is preferably withdrawn from separator 24 in a split stream, a portion being sent through regenerated catalyst recycle line 30, and another portion through regenerated catalyst line 3| leading to the conversion or reaction system. Both catalyst outlet lines 30 and 3| are pressure standpipes similar to standpipe H in that they are provided with means for introducing an aerating medium at suitable points along their length so as to maintain the catalyst flowing therethrough in a condition wherein it has the flow characteristics of a liquid, such means being lines 32 leading into recycle line 30 and lines 33 leading to catalyst outlet line 3|. The quantity of catalyst withdrawn and recycled through line 30 is preferably maintained within predetermined limits so as to maintain the temperature in regeneration zone 2| within required limits in accordance with the principles of operation described in U. S. Patent 2,253,486 to Arnold Belchetz, dated August 19, 1941.

Regenerated recycle catalyst is fed from the bottom of standpipe 30 through a suitable feeding means such as a slide valve 34 into an inlet line 35 leading to a heat exchanger or catalyst recycle cooler 22. Regenerated catalyst thus introduced is picked up by air introduced into line 35 through line 36, the quantity of air thus introduced being sufiicient together with any air introduced through line 20' to efiect combustion to the required extent in the regenerator 2|.

From line 35 the mixture of air and regenerated catalyst passes through exchanger or cooler 22 wherein the temperature of the recycled catalyst is lowered by indirect heat exchange with a cooling medium circulated through the exchanger by lines 31 and 38.

Regenerated catalyst is fed to the conversion system from standpipe 3| through a suitable feeding means at the bottom thereof such as a slide valve 39 and line I into the stream of vapors to be converted passing through line 2 as previously described.

To illustrate the operating conditions preferably maintained in the regeneration zone, reference is made to the data tabulated in the following Tables 2A and 2B wherein regeneration zone operating conditionsare shown corresponding to the conversion runs tabulated in Tables 1A and 1-B.

TABLE 2 Spent catalyst, lbs./hr 632,840 Cooled recycled catalyst, lbs/hr 1,750,000

Ratio by weight recycled/spent 2.77 Inlet temperature, spent catalyst, F 900 Inlet temperature, recycled catalyst, F. 840 Temperature, mixture recycled and spent catalyst, F 850 Temperature, regeneration chamber, F 1,000 Regeneration dimensions:

(a) Height, ft 50 (b) Diameter, ft 18 Regeneration velocity:

(a) Base 1.62 (12) Top 2.59 Air feed, lbs./hr 88,350

Catalyst concentration, lbs/cu. ft.:

erably a series of such separators such as cyclone (a) Regenerator 2 parators or the like 41a, 41b and 41c. In each (1)) Outlet line 1- Of these a part of the suspended particles i sepa- Weight per cent of coke based on oil rated and withdrawn from the bottom of the sepfeed 4.85 arators through tail pipes 48a, 48b and 480. Ma- Coke per cent by weight on spent terial discharged from these tail pipes may be catalyst y d to a regenerated catalyst recovery hop- Carbon per cent by weight on regenper 49 through lines 50a, 50b and 500 by way of erated catalyst 0!? transfer line 5| by means or a suitable fluid n- Catalyst contact time, seconds 352 veying medium, such as steam, supplied by jets Pressure in regenerator, lbs./sq. in.: through lines 52a, 52b and 520.

(a) Inlet cone 16 From separator 4111 the suspension passes by (b) Outlet cone.. 9 'line 52 to a Cottren precipitator 5: r other t- TABLE 2--B Run Number 5254 5325.4 5320.1 5205 5330.1 5330.5

Spent Catalyst, LbsJHr 00 00 105 Temperature, Regeneration Chamber, F 1,050 1,030 1,045 1,010 1,010 1,025 Regenerator Dimensions: V

(0) Height, Fr 12 12 12 24 12 12 (0) Diameter, Tn 4 4 4 3 4 4 Regenerator Gasvelocity:

Ram 1.52 1.38 1.73 1.62 1.34 1.30 5 Top 1. 01 1.51 1.13 1.73 1.30 1.42 Catal t Concentration, Lbs/Cu. Ft:

5 mm 0.19 0.21 0.24 0.28 0.10 0.30 (b) P amr 12.4 12.4 10.0 9.6 11.2 12.1 Wt. Percent of Coke Based on Oil Feed 5. 1 3. 5 4.4 1.9 3. 5 4.4 Coke Percent by Wt. on Spent Catalyst 1.63 1. 57 1.78 1.41 1. 73 1.35 Carbon Percent by Wt. on Regenerated Catalyst 0.40 0. 47 0. 58 0.53 0.47 0.38 Catalyst Contact Time. Seconds 530 530 .366 540 540 350 Pressure in Regeneratoi', Lbs/Sq. in 11.0 12.6 12.5 12.5 13.5 13.5 I

A preferred system for the recovery of residual catalyst present in the overhead product from separator 24 is illustrated in Fig. 1. Referring to this figure gaseous regeneration products mixed with a relatively small portion of the catalyst originally present therein are withdrawn from the upper part of hopper 24 through line 25 and pass to a series of separating zones constituted by suitable gas-solids separators such as cyclones, Cottrell precipitators, filters, or the like, the recovered material being eventually returned to hopper 24 through line 26. a

The gaseous suspension in outlet pipe 25 consists essentially of flue gas and residual suspended regenerated catalyst including both relatively fine and relatively coarse particles. This suspension is preferably supplied to the recovery system at a superatmospheric pressure sufliciently high to impel it completely therethrough and into the atmospher from the final separating zone, the pressure in the successive separating zones being progressively lower in the direction of the flow of the suspension by reason of the pressure drop in the interconnecting linesand gas-solids separating means. Incidental to such separation it has been ascertained that classification and segregation of the particles occurs, particularly segregation of the extremely fine particles in the final separating zone, this zone being a Cottrell precipitator as shown, or any other suitable means utilized for the separation of the last increment of the suspended particles.

In certain instances cooling of the gaseous suspension passing through line 25 may be desirable, as for example by passing it through a suitable cooler or heat exchanger through which a heat exchange medium is circulated through lines 46 and 41, thereby effecting a reduction in temperature and volume of the suspension passing therethrough, it being understood however that such cooling is not essential and may be omitted.

Fro'm cooler 45 the gaseous suspension passes by line 46 to a suitable gas-solids separator, or pretparticles being collected in the bottom hopper 54 of the precipitator, and the separated gas exiting overhead to the atmosphere through line 55.

Cottrell precipitator 53 is preferably operated under approximately atmospheric pressure, and a pressure reduction valve may be provided in line 52 for regulating the pressure so that the desired pressure may be maintained in separator 53 irrespective of the pressure in the discharge line leading from cyclone 4111.

A continuous stream of previously separated relatively coarse particles preferably is supplied to a mixing zone in hopper 54 through line 56 from hopper 24 for the purpose oi mixing with and bringing the fines to a condition which may be described as flowable. Coarse particles for this purpose supplied through line 56 may be supplemented by fresh or make-up" catalyst supplied to the system from hopper 51 through line 58. Such make-up" catalyst is necessary to replace unavoidable losses of catalyst from the system and to compensate for any gradual permanent decrease in activity of the circulated cataiyst. An aerating medium preferably is supplied to the bottom of the hopper 54 by means not shown, but similar to elements 21 and 28, for aerating the mixture therein and maintaining it in a readily fiowable condition.

From hopper 54 the aerated mixture flows into an outlet standpipe 59, operating in a manner similar to standpipe 30, an aerating fluid such as steam being supplied thereto along its length through lines 60. standpipe 59 is preferably of a height suilicient to provide a fluid head to balance the pressure diiferentlal between the hopper 54 and the zone of relatively high pressure to which the separated fines are returned to the system, for example hopper 49 as shown or hopper 24. From standpipe 59 the mixture is discharged by means of a suitable valve such as a slide valve 6| into transfer line ll. Steam or other suitable 9 conveying medium is supplied by line 62 to line ii to convey the mixture to hopper 49 or if desired to any other zone.

In passing through line the mixture is mbined with the streams from lines 50a, 58b and 50c and passes to collecting hopper 49 wherein the combined fractions are separated in the bottom of the hopperyand the gaseous suspending medium is separated overhead through line 83 and passes into line 46 leading to the cyclone separators. From hopper 49 the separated solids are withdrawn through a standpipe 64 operating on a principle similar to standpipes 38 and 59 and to which a suitable aerating medium is supplied through lines 65. standpipe 64 is preferably of a height sufficient to balance the diflerence in pressure between hopper 48 and hopper 24. From standpipe 64 the solids are fed through a slide valve 66 into transfer line 26 wherein they are suspended by suitable conveying fluid such as steam supplied through line 61 and conveyed therethrough to hopper 24 and combined with the initially separated material.

The operation of the regenerated catalyst recovery system may be further exemplified by reference to conditions obtaining in the specific regeneration operation illustrated by Table 2-A. Pursuant to this example, the proportions of regenerated powdered cracking catalyst separated in the various separating zones, based upon the quantity of material entering the system through line l9, were approximately as follows:

Per cent Hopper 24 75 Cyclone 41a 20 Cyclone 41b 2.5 Cyclone 41c 1 Precipitator 53 1.5

The percentage recovery in hopper 24 based upon the total solids passing through outlet 23 including the recycled catalyst from line 30, is substantially greater than on the above basis, amounting to about 95%.

In this example, pressure conditions obtaining throughout the regeneration system were approximately as follows:

Pressure,

Zone Lbs/Sq. In.

Regenerator 21 (base) Hopper 24. Line 46 Cyclone 47a (inlet) Cyclone 47b (inlet) Cyclone 47c (inlet)- Prccipitator 53 (inlet) i; alve 61 Hopper 49. Valve 66.

Referring back to the conversion side of the system, the vaporous conversion products exiting overhead from hopper 9 through line l3 areof cyclone separators similar in design and operation to those previously described with respect to the regenerated catalyst recovery system. The reaction vapors withdrawn through line l3 preferably contain a relatively small proportion of the total catalyst introduced through line 2, since hopper 9 is of a design capable of separating about 90% or more of the total catalyst.

The vapor mixture in linev l3 may be cooled, if desired, by passingit'through a cooler 18, to a temperature above the condensation point of the vapors, but sufl'lciently low to substantially reduce the volume of gas passing to the cyclone separators, thereby increasing the concentration of the catalyst in the vapors and thus facilitating catalyst separation. This cooler may be omitted and the suspension passed directly by by-pass line H and line I2 to a series of cyclone separtors 13a, 13b and 13c. In each of the cyclones a portion of the catalyst is separated in the bottom of the cyclone and the separated gas withdrawn overhead and passed to the subsequent separator stage. The separated fractions are picked up from the discharge end of the tail pipes 14a, 14b and He by a jet of steam introduced through valves 15a, 15b and 150, respectively, and conveyed to a collecting drum 16 through transfer line 11. Due to pressure drop in the cyclone recovery system and interconnecting lines, drum [6 is at a pressure substantially lower than the pressure prevailing in hopper 9. The recovered catalyst is withdrawn from drum 16 through a catalyst standpipe 18 of a height suflicient to balance this difference in pressure, an aerating medium being supplied to the standpipe through lines 19.

Catalyst is discharged from the bottom of standpipe 18 through valve 80 and is conveyed by a suitable medium such as steam introduced through line 8| to hopper 9 by line l4.

Normally a very small proportion of the catalyst is left in the vapors exiting from the last cyclone 13c and this residual catalyst is preferably recovered by partial condensation and scrubbing in a lower section of the fractionation tower, as described in U. S. Patent No. 2,374,073 to Arnold Belchetz, dated April 17, 1945. This fraction will normally constitute less than 1%, for example about 0.2% of the total circulated catalyst.

The fractionator for the recovery and separation of the conversion products may be of conventional design, that shown being illustrative of a suitable type utilized in the commercial operation for which data is given in Table 1A. Inspections'of the feed and products for this particular operation are tabulated in Table 3.

From cyclone 130 the mixture of conversion product vapors and residual catalyst is passed by line 82 to the base of a combination scrubbing and fractionating tower 83. The mixture passes upwardly in the tower over baifies 84 in countercurrent flow with liquid reflux or condensate supplied through lines 85 and 86. A partial condensate consisting of the highest boiling constituents of'the vapors is thus formed in which the residual catalyst is concentrated and withdrawn from the bottom of the tower through line 81. A portion of the bottoms product is diverted through line 88 through a cooler 89 and returned by line 85 over the bailles to provide the required temperature gradient in the scrubbing section. A heavy gas oil fraction may be withdrawn by side draw-off line 98 and passed by pump Sithrough cooler 92. A portion of this gas oil may be withdrawn from the system through line '93, and another part 11 returned to the tower for refluxing through line 8'.

A light gas oil may be withdrawn through side draw-oil line 94 and stripped in side stripper 99, the overhead product being returned through line 96 to the tower, and a light gas oil being withdrawn from the system through line 91.

Condensate may be withdrawn from the top portion of the tower through line 91, passed by pump 99 through heat exchanger 99, and returned through line I 09 at a suitable temperature to provide the required reflux at the top of the tower.

In the operation shown naphtha and lighter products are withdrawn overhead and may be further separated in a suitable recovery system including an absorber, stripping tower, and stabilizer, or by other suitable means, to produce the lighter products indicated on the following Table 3.

- TABLE #3 do to coke.

Products API Gravity Yields Bbla. per day The high boiling fraction containing the resid super-heat coil I03 in furnace I, or preferably by heating and vaporizing this fraction in a separate coil and then combining the vapors with those from the super-heat coil. The preferred feed preparation means utilized to vaporize and heat the hydrocarbon feed is dependent upon the character of the charging stock. In certain cases as I with a gas oil feed of the type given in Table 3 the preheated liquid feed may be introduced directly through line I02 to coil I03 which in this instance vaporizes and super-heats the feed, steam preferably being added to the coil in an amount as indicated in Table lA. With heavier feed stocks the modified form of feed preparation means shown, including a vaporizing furnace I96 and a flash drum I99, is suitably employed. In the latter case preheated liquid feed stock from line I94 is fed by pump I05 into vaporizing coil J" in furnace I96. Steam is introduced from line Ill in suitable quantities to coil I96. The heated charge is then flash evaporated in flash drum I99, p the feed vapors being taken overhead through line I02, and bottoms withdrawn from the system through line I I I. Additional steam may be added to the feed vapors in line 3 through line 9.

In the foregoing a preferred procedure and arrangement of apparatus for practicing the invention has been described, however, it will be apparent to those skilled in the art that various modifications thereof may be made without 'departing from the essential features of the invention.. For example, the movement and circulation of the catalyst to and from the various zones may be efifected by solids pumps" such as described in Kinyon Patent 1,553,539 in place of catalyst standpipes. Also, other types of gas-solids separating means for recovering finely divided solids from the gaseous products withdrawn from the conversion and regeneration zones may be sub- 12 stituted for those shown. The fresh or regenerated catalyst need not be introduced into the conversion zone in mixture with the feed vapors, but may be introduced separately as for example by providing a separate line interconnecting the bottom of standpipe 3| and reactor 6 and conveying the catalyst therethrough by a suitable medium such as ste'am,'or standpipe II may terminate directly in the base of reactor 9. By a further modification the withdrawal of used catalyst as well as the addition of spent catalyst to the reactor may be eifected through inlets and outlets separate from those utilized for the introduction and withdrawal of vapors, as for example by a process flow such as that shown in Fig. 3. Certain variable operating conditions in the practice of the process may follow and be controlled pursuant to conventional practice in the art of vapor phase catalytic cracking of high boiling hydrocarbons, including such factors as the selection of suitable charging stock, catalytic material, conversion temperatures, and pressures.

Primary operating variables distinctively controlled, pursunt to the present invention, are the" rate of feed of the high boiling hydrocarbons and the rate of feed of the catalyst to the reaction zone and the weight of catalyst in said zone.

The rate of feed of the hydrocarbon charge to a reactor of given dimensions is maintained within such limits that the upward velocity of the vapors through the zones is relatively low and sufflcient to form a dense phase or mass of catalyst therein. Conversely, for the conversion 01' a given quantity of charging stock in a given unit of time, the cross-sectional area of the reactor must be of the dimensions requires to provide the desired low vapor velocity therein.

The rate of fresh catalyst feed is dependent upon the desired average catalytic activity of the dense phase of catalyst in the conversion zone, and fresh catalyst is continually added at a rate adapted to maintain such activity at the desired value as the conversion proceeds. Used catalyst is withdrawn at the same average rate as fresh catalyst is added, therefore, the average time a catalyst particle remains in the reactor (catalyst resident time) is determined by the catalyst feed rate and may be calculated by dividing the weight of catalyst in the reactor by the catalyst feed rate per minute. The concentration (density) of catalyst in the dense phase is dependent primarily upon the particular low vapor velocity maintained. Within limits, an increase or decrease of the rate of catalyst feed apparently has no substantial or significant eflect on the concentration (density) of the dense phase. A further factor effecting the optimum feed rate of catalyst to the conversion zone is a condition which may be termed the catalyst level phenomenon. Dependent upon the height of the reactor and with low velocities, two distinct phases of catalyst concentration are present in the reactor, a lower dense phase and an upper phase wherein the concentration of catalyst is relatively very dilute, the boundary or interface between the phases being at a horizontal plane intermediate the vapor inlet and outlet. It has been ascertained that the distance of this interface from the top outlet 1 in reactor 6, other conditions being fixed, is dependent upon the rate of catalyst feed and that the magnitude of this distance varies inversely with the rate of catalyst feed. Accordingly, the rate of feed is normally and preferably regulated (or, a fixed catalyst feed rate being assumed, the height of the reactor fixed) so that the upper 13 level of the dense ca ly phase will be confined to the upper part of the reactor, this level or interface being indicated in reactor 6 by dotted line A and in regenerator 2| by dotted line B.

The weight of catalyst in thereactor (a fixed height of reactor being assumed) is dependent upon the concentration of the dense phase (in turn dependent primarily upon the particular low velocity maintained) and the distance of the upper level of the dense phase from the upper vapor outlet (the latter in turn being dependent upon the catalyst feed rate).

Accordingly, proceeding on the above noted principles pursuant to this invention, the rate of feed of hydrocarbon vapors and rateof feed of the catalyst to the conversion zone are preferably controlled within such limits that the vapors of the high boiling hydrocarbons flow upwardly through the conversion zone at a velocity suiliciently low to form a dense phase or mass of the catalyst in said zone, and fresh catalyst is added to said dense phase and corresponding amounts boiling p int of 470 F. and end point of 736 F.

(A. S. '1. M.). The catalyst employed was a sill-- ca-alumina type of cracking catalyst consisting of powdered Super-Filtrol, the particle size distribution of the catalyst being approximately 10% -10 microns, 20% -20 microns, 2030 microns, 20% -40 microns, and 30% microns. In Table 4, the ratio of the weight of catalyst feed per hour to the weight of oil feed per hour constitutes the cat/oil" weight ratio. Catalyst resident time cat. time mins." for both reactor and regenerator was calculated by dividing the weight of catalyst in the reactor by the catalyst feed rate per minute. The vapor velocities at the inlet of the reactor were calculated on the assumed basis that the catalyst volume was negligible. The outlet vapor velocity was calculated on the basis of the outlet conditions of temperature and pressure, and mols of product produced. The arithmetical average of inlet and outlet velocities was divided by the reactor length to obtain the superficlal oil contact time (011 time) TABLE IV 35 .4. P. I. Mid-Continent gas oilCatal1/st Super-Filtrol Reactor-Vertical 2% In. x 12 of used catalyst withdrawn therefrom at a rate adapted to maintain the average catalytic activity of said mass of catalyst at a suitable value and preferably also at a suitable rate to maintain the level of said dense phase within the confines of the upper part of the reaction zone.

The eilect of variations in the several operating conditions over a wide range is disclosed by the illustrative runs given in the following Table 4. In these runs the feed stock consisted of a 35 910 12 9 l. 3 39. 4 3. 7 3. 6 l4. 1 910 8. 4 3. 1 43. 3 2. 3 4. 8 7. 8 910 6. 9 2. 5 37. 5 3. 5 10. 4 0. 0 910 14. 3 2. 7 30. 6 1. 6 5. 4 0. 0 950 14. 5 3. 7 30. 4 1. 1 5. 2 0. 0 ,M 23.8 2.0 32.4 1.1 4.5 0.0 905 34.4 1.4 26.2 1.2 4.1 4.2 910 8. 3 3. 8 37. 4 1. 9 6. 4 1. 8 905 4. 6 3. 3 48. 8 4. 0 5. 9 13. 3 905 3. 5 5.4 48. 2 3. 2 7.1 3. 4 900 162. 0 0. 1 23. 4 2. 9 3. 7 9. 5 900 15.7 1.0 38.2 3.8 4.9 7.0 895 8. 8 1. 5 44. 3 4. 6 5. 5 7. 1 900 10. 0 1. 8 43. 2 3. 4 5. 3 7. 0 900 9. 3 1. 7 44. 7 3. 9 5. 4 7. 1 900 6. 4 2. 0 52. 2 4. 7 6. 4 9. 9 ill) 6.2 2.0 48.4 5.0 6.4 10.0 895 41. 3 0. 3 29. 4 4. 8 3. 6 10. 2 895 4. 4 2. 1 55. 4 6. 7 10. 6 0. 0 900 5. 8 l. 6 56. 8 6. 7 6. 3 10. 1 895 5. 8 1. 5 54. 7 6. 7 6. 3 10. 0 900 5. 7 1. 6 54. 4 6. 6 6. 5 10. 0 910 5. 6 1. 7 57. 1 6. 3 6. 0 10. 1 900 4. 8 1. 8 62. 4 7. l 6. 8 10. 1 905 5.4 3.6 50.9 3.0 8.0 9.7 895 7. 9 3. 5 44. 2 2. 7 6. 5 9. 1

ReactorVertical 2% In. x 12 Ft.

* Ci 5 A: a u u u g o .1 Velocity o g m g g g m 5 0 53 .2 0o 6 5, O 8 L. o O -----5 h m a an E 9 .6 F? a e a 53 I? .s 0 B1 2' 5 g i o '5 O 3 .s Q .s O B H O O B 3 Ft. Tubew.lhr./w. Greater Than 3.0

0.44 2. 3. 7. 0 36. 1 60. 6 4. 8 2. 9 1. 5 0. 44 1. 96 3. 16 10'. 6 38. 8 56. 7 5.6 3. 6 1. 8 0. 35 1. 1 2. 2 13. 2 33. 1 62. 5 4. 7 3. 1 2. 1 0.40 1.6 3.0 10. 3 29.0 69.4 2.7 2. 1 1.3 0. 36 1. 6 3. 1 l0. 3 28. 8 69. 6 2. 7 2 3 1.1 0.34 1.8 5.6 5.5 26.7 67.6 3.3 5.2 1.0 0.27 2.4 3.5 4.3 26. 1 73.8 1.5 1. 6 0.9 0.30 1.3 2.5 10.9 34.5 62.6 3. 9 3.1 1. 5 0.28 1.6 2.5 10.3 41.1 51.2 7.6 4.6 2.4 0. 28 1. 87 1. 6 13. 4 42. 9 51. 8 5. 4 3. 6 2. 8 0. 50 2. 07 3. 58 79 22. 9 76. 6 0.6 1. 2 1. 4 0. 67 2. 0 2. 9 5. 8 34. 3 61. 8 4. 7 2. 5 2. 3 1. 23 1. 8 2. 9 10. 0 38. 3 55. 7 5. 3 3. 1 3. 7 0. 79 1. 8 2. 7 8. 8 38. 3 56. 8 5. 9 3. 0 2. 4 0. 75 1. 8 2. 7 9. 4 39. 2 55. 3 5. 6 3. 4 2. 9 0. 48 1. 5 2. 2 l0. 0' 44. 8 47. 8 7. 1 4. 2 3. 1 0.55 1.5 2.3 10.3 40.7 51.6 7.8 4.0 3.2 0. 81 3. 0 3. 7 3. 0 26. 2 70. 6 3. 3 2. 2 2. 0 1. 15 O. 7 1. 5 14.6 44. 9 44. 6 7. 8 5. 1 5. 6 0. 96 1.5 2.3 10.9 46.7 43. 2 8.0 5.4 4.6 0. 85 1. 5 2. 3 10. 9 45. 0 45. 3 8. 8 5. 1 4. 2 0. 65 1. 5 2. 2 11. 1 42. 4 45. 6 9. 1 5. 8 4. 6 0. 73 l. 5 2. 5 11. 5 46. 7 42.9 9. 4 5. 8 3. 7 0. 75 l. 3 2. 3 13. 4 50. 5 37. 6 9. 9 5. 9 5. 0 0. 74 1. 2 1. 8 12. 1 45. 5 49. 1 5. 8 4. 5 2. 7 0. 37 1. 5 2. 2 10. 9 42. 3 55. 8 1. 4 4. 0 2. 0

Tubew./hr./w. Less Than 3.0

0. 47 0. 67 1.15 17. 0 49. 7 31. 8 l2. 3 7. 6 .7. 6 0. 70 l. 1 1. 6 15. 8 46. 9 34. 1 12. 2 6. 5 9. 0 0. 52 1. 1 1. 7 15. 8 47. 5 34. 4 12.0 7. 7 7. 4 0.39 2.6 3.3 11.2 45.7 40.2 11.4 6.6 4.5 0. 56 0. 43 0. 16. 7 47. 0 35. 3 9. 3 8. 1 8. 0 0. 74 1.1 1. 7 14. 6 45. 3 40. 8 10. 0 6. 0 5. 6 1. 84 0. 8 1. 3 21. 9 45. 4 34. 1 10. 3 7. 2 10. 7 0. 59 0. 7 1. 1 15. 2 51. 6 27. 5 13. 3 8.1 9. 2 0. 50 0. 7 1.1 16. 4 53. 3 27. 5 12. 0 7. 5 9. 1 0. 53 0. 45 0. 7 17. 6 .50. 6 18. 0 l6. 0 9. 1 16. 6 0. 53 0. 43 0. 71 17. 0 43. 8 19. 9 l6. 5 13. 8 15. 7

From the specific examples given it will be evident that an operating condition of primary importance in the practice of the invention is the relatively low velocity and resulting high concentration of catalyst maintained in the conversion and regeneration stages. Referring for example, to Table 1A it will be noted that the catalyst concentration in reactor 6 was 18 lbs. per cu. ft. as compared with a catalyst concentration of about 1 lb. per cu. ft. in the reactor inlet line 2.

A. P. I. Mid-Continent gas oil having an initial 76 Under such high concentration conditions, the

fiow of the catalyst may be described as that of a compacted turbulent mass or moving bed.

In order to provide this dense or concentrated phase of the catalyst in the conversion and regeneration stages, low vapor velocities of the or der of about 6 ft. per second or less, and preferably of about 4 ft. per second or less are contemplated. Vapor velocity ranges regarded as especially suitable for the practice of the process in vertical reaction vessels of the type shown in Fig.

2, are about 6 to 0.5 ft. per second, preferably about 1.5 to 2 ft. per second. Even lower vapor velocities than these indicated minimums may be utilized to advantage with the modified type of conversion system illustrated in Fig. 3. It is further contemplated that the practice of the process, under most conditions, may be satisfactorily effected, utilizing a catalyst to oil feed weight ratio within the range of about. 0.5:1.0 to

:1.0 and preferably within the more restricted range of about 2:1 to 8:1, and with a value of w./hr./w. within the range of about 1.0 to 25.0

and preferably within the more restricted range of about 2.5 to 10.0.

The modified process fiow illustrated in Fig. 3

differs from that of Figs. 1 and 2 in that the catalyst is introduced to and withdrawn from the conversion zone and regeneration zone by catalyst inlets and outlets separate from the inlet and outlet for all or the major proportion of the vapors undergoing conversion, or the regeneration fluid. Elements of Fig. 3 having a generally similar function to those described in detail in connection with Figs. 1 and 2 are designated in Fig.

3 with a similar numeral with the suifix a and hence detailed description of these elements is superfluous. Operating variables likewise may be controlled for this process flow pursuant to the principles disclosed in connection with Figs. 1. and 2. Feed vapors are introduced through line 3a and a suitable manifold 51: substantially uniformly throughout the bottom area of the reactor so,that the dense phase of catalyst therein is maintained in an aerated flowable condition.

1 Regenerated catalyst'from standpipe 39a is introduced by a current of a suitable conveying medium such as steam (or a mixture of steam and I part of the feed vapors supplied through line H 4) through catalyst inlet line I a and flows laterally through the reactor transversely to the entering vapors and is withdrawn at the opposite side of the reactor at a controlled rate through standpipe Ila by means of valve I 8a. The height of level of the dense phase of catalyst in the reactor 6a is regulated by operation of valve l 8a. and in this I respect differs .from the flow described in connection with Figure 2. The horizontal level of H2 and returned by line H3 to the reactor at the steam stripping zone above the catalyst standpipe outlet II a. Similarly in the regeneration zone, the spent catalyst is introduced thereto by a separate outlet l9a from standpipe Ha and withdrawn at the opposite end of the regeneration zone. by a separate outlet standpipe 30a. The level of catalyst in the regeneration zone is controlled in the same manner as in the conversion zone so that only a relatively small amount of catalyst passes out of the zone overhead. through the outlets 23a to the gas-solids recovery system. The vapor velocities used in this type of process flow may be quite low, the limiting quantity being that required to maintain the body of catalyst in an aerated or flowable condition.

Any of the various known types of cracking catalysts may be utilized in the practice of the invention. The preferred catalysts are those of the silica-alumina, or silica-magnesia type adapted to produce a satisfactory yield of high octane gasoline. Either silica-alumina catalyst consisting of activated clay prepared by the acid treatment of natural clays, for example the commercial product fSuper-Filtrol or a synthetically prepared silica-alumina catalyst such as those disclosed in U. S. Patents Nos. 2,391,481 and 2,391,482 to Robert Ruthruif, both dated December 25, 1945, may be employed. The catalyst is preferably employed in finely divided or powdered condition, for example with particles ranging from about 1 to microns. Other conditions such as temperature, pressures, feed stocks, and the like, may be selected and controlled pursuant to conventional practice in the art of vapor phase catalytic cracking of high boiling hydrocarbons.

We claim:

1. In a catalytic hydrocarbon conversion system wherein hydrocarbon vapors pass upwardly through a reaction zone in contact with a dense turbulent suspended catalyst phase and thence to an enlarged catalyst settling zone wherein spent catalyst from said settling zone is then zones and settling zones, and positively injecting said fines by means of steam into the settled catalyst flowing from said respective settling zones.

2. In a catalytic hydrocarbon conversion system wherein hydrocarbon vapors pass upwardly through a reaction zone in contact with a dense turbulent suspended catalyst phase and thence to an enlarged catalyst settling zone wherein spent catalyst from said settling zone is then transferred to a regeneration zone and contacted with regeneration gases flowing upwardly therein at such a velocity as to maintain the catalyst in a dense turbulent suspended catalyst phase, wherein gases from the regeneration zone pass to an enlarged settling zone from which catalyst is transferred to the conversion zone and wherein catalyst fines are separated from gases and vapors leaving said respective settling zones, the method of returning said catalyst finesto the system which comprises discharging said fines to a hopper exterior of the contacting zones and settling zones, and positively injecting said fines by means of steam into the settled catalyst flowing from a settling zone.

3. In a catalytic conversion system of the type wherein a gas or vapor passes upwardly in a contacting zone and is contacted in said zone with a dense turbulent suspended catalyst phase, wherein the bulk of the catalyst is removed from gases and vapors in an enlarged settling zone, wherein a body of settled catalyst is maintained in the lower part 01 the settling zone, and wherein residual catalyst material is removed from gases or vapors by means of cyclone separators, the method of combining cyclone separated catalyst with settled catalyst which comprises discharging said cyclone separated catalyst into an external collection hopper and positively impelling cyclone separated catalyst into the bed of settled catalyst in said settling zone.

PERCIVAL C. KEITH. JOSEPH W. JEWELL.

REFERENCES CITED UNITED STATES PATENTS Name Date Miller Apr. 7, 1931 Number Number Number Name Date Osterstrom et al. Aug. 23, 1932 Odell Dec. 18, 1934 Subkow Feb. 11, 1941 Huppke Feb. 11, 1941 Belchetz Aug. 19, 1941 Becker et a1 Jan. 27, 1942 Voorhees Feb. 17, 1942 Hemminger Oct. 27, 1942 Hemminger Nov. 24, 1942 Degnen Dec. 15, 1942 Voorhees Feb. 9, 1943 Guild Feb. 23, 1943 Stratford et al May 18, 1943 Thiele et al Aug. 10, 1943 Payne July 11, 1944 Scheineman July 11, 1944 FOREIGN PATENTS Country Date Australia Sept. 20, .1937 Germany Sept. 8, 1931 

1. IN A CATALYTIC HYDROCARBON CONVERSION SYSTEM WHEREIN HYDROCARBON VAPORS PASS UPWARDLY THROUGH A REACTION ZONE IN CONTACT WITH A DENSE TURBULENT SUSPENDED CATALYST PHASE AND THENCE TO AN ENLARGED CATALYST SETTLING ZONE WHEREIN SPENT CATALYST FROM SAID SETTLING ZONE IS THEN TRANSFERRED TO A REGENERATION ZONE AND CONTACTED WITH REGENERATION GASES FLOWING UPWARDLY THEREIN AT SUCH A VELOCITY AS TO MAINTAIN THE CATALYST IN A DENSE TURBULENT SUSPENDED CATALYST PHASE, WHEREIN GASES FROM THE REGENERATION ZONE PASS TO AN ENLARGED SETTLING ZONE FROM WHICH CATALYST IS TRANSFERRED TO THE CONVERSION ZONE AND WHEREIN CATALYST FINES ARE SEPARATED FROM GASES AND 